MoS2 CATALYST SYSTEM FOR THE CONVERSION OF SUGAR ALCOHOL TO HYDROCARBONS

ABSTRACT

Cellulose and hemicellulose from biomass can be broken down to C6 and C5 sugars and further converted to corresponding sugar alcohols. It is now found that a new catalyst, MoS2, is active for the hydrogenation of sugar alcohols to hydrocarbons. Combining the technologies listed above allows us to convert the cellulose/hemicellulose to liquid hydrocarbons.

CROSS REFERENCE TO RELATED APPLICATIONS

This is a divisional application that claims benefit under 35 USC §120to non-provisional U.S. application Ser. No. 13/233,256 filed Sep. 15,2011, titled “MoS2 CATALYST FOR THE CONVERSION OF SUGAR ALCOHOL TOHYDROCARBONS,” and claims benefit under 35 USC §119(e) to U.S.Provisional Application Ser. No. 61/444,004 filed Feb. 17, 2011,entitled “MoS₂ CATALYST FOR THE CONVERSION OF SUGAR ALCOHOL TOHYDROCARBONS,” both of which are incorporated herein in their entirety.

STATEMENT OF FEDERALLY SPONSORED RESEARCH

None.

FIELD OF THE DISCLOSURE

The present invention relates generally to catalysts that convertcarbohydrates to fuel range hydrocarbons.

BACKGROUND OF THE DISCLOSURE

Methanation is an important process for upgrading coal and biologicalmaterials to useful fuel gases. In coal gasification, methanation is thecatalytic conversion of hydrogen and carbon monoxide to methane.Methanation is also used to produce biomethane from organic sources ofenergy. One method of distributing vast quantities of coal and biomassenergy is to gasify the coal to produce synthesis gas (syngas) and thenconvert the syngas to substitute natural gas (SNG) via methanation(Reid, 1973). Current methanation processes use a nickel (Ni) catalystwhich imposes certain operating limitations (FIG. 1) because of itssusceptibility to deactivation by surface carbon, high temperaturerequirements, and poisoning by various sulfur compounds. The stringentprocess restrictions shown in FIG. 1 require additional steps forsuccessful use of nickel catalysts (Walston, 2007).

A major restriction for nickel catalysts comes from their extremesensitivity to poisoning by sulfur compounds which are always present incoal-derived synthesis gas. Syngas processed by nickel catalysts must bepurified to below 20 ppb sulfur to avoid poisoning of the catalyst eventhough pipeline natural gas can contain up to 4 ppm hydrogen sulfide.Nickel catalysts can also be irreversibly poisoned by carbon fouling. Toavoid carbon fouling the H₂/CO ratio has to be adjusted to valuesgreater than three by the water-gas-shift (WGS) reaction. Nickelcatalysts are also deactivated by sintering at high temperatures (>450°C.). The methanation reaction is so highly exothermic that a 5 molepercentage reduction in carbon monoxide (CO) concentration due to themethanation reaction results in about a 260° C. (500° F.) increase inreactor temperature. Ni catalysts are so active that it is hard toobtain less than 100% CO conversion even at high space velocities.Therefore, in industrial plants, around 90% of the product gas from themethanation reactor is recycled back to dilute the concentration of COto less than 5% (mol) in the feed gas. Lurgi designed the Great PlainsSynfuels Plant in North Dakota using this conventional methanationtechnology (Lukes, 2003; Anand, 2007). This is the only existing exampleof a commercial coal to SNG facility and has been operating since July1984.

Improvements to the conventional methanation process were made by HaldorTopsoe and Johnson Matthey by developing a high temperature methanationtechnology. Haldor Topsoe developed a Ni-based, thermally stablemethanation catalyst, MCR-2X, for a high temperature methanationprocess, trade marked as TREMP™ (Total REcycle Methanation Process)(Udengaard, 2006). This catalyst can be operated at higher temperaturesthan previous Ni catalysts and does not sinter below 700° C. (˜1250°F.). The methanation unit is made up of a series of adiabaticmethanators with inter-stage cooling and gas recycle to control thereactor temperature. As a result of the higher temperature tolerance ofthe catalyst, per pass CO conversion can be increased. This allows forsmaller recycle ratio and methanation reactor size resulting in lowerCapEX and OpEX than conventional methanation. The higher effluent gastemperature at the reactor outlet can be used to generate superheated,high pressure steam to further improve energy efficiency of the process.This process was demonstrated in a single 1.7 mmscf/d reactor in the1980's.

Johnson Matthey has developed a similar high temperature methanationtechnology. Their catalyst is a modified version of their Ni-basedpre-reformer catalyst traditionally used in hydrogen and/or ammoniaplants. A newer generation of high temperature methanation catalyst,CRG-LH, was formulated in the 1990's, for improved thermal stability.According to Johnson Matthey, this catalyst has been tested on a pilotscale reactor (¾″ diameter, 8 ft long) at 500-600° C. (˜1100° F.), 30-50bar for over 1000 hours. The results showed acceptable thermalstability. Demonstration of this catalyst on a larger reactor seemsnecessary to fully evaluate its commercial capability.

Other improvements to the conventional methanation process are combinedshift/Methanation (Graboski, 1975), sulfur-tolerant methanation andcatalytic steam gasification. Combined shift/methanation technologiessuch as Conoco's SUPER-METH™ (Kock, 1979; Sudbury, 1980) Parson'sRMPROCESS™ (Dissinger, 1980; White, 1975; White, 1976), United Catalyst,ICI, and UOP utilize water formed in methanation for water-gas-shift andhence combine the water-gas-shift and methanation reactions into asingle reactor. Apart from an acid gas removal unit upstream of themethanation reactor for H₂S removal, an additional acid gas removal unitis required downstream of the shift/methanation process to remove CO2.These technologies were piloted in the 1970s but have not proven to becommercially viable due to costs, complexity, scalability or othercomplications associated with the demands of refinery methanationprocesses.

The sulfur-tolerant methanation process developed by the Gas ResearchInstitute (GRI) (U.S. Pat. No. 4,491,639; EP0120590; Happel, 1979;Happel, 1981; Happel, 1982; Happel, 1983; Happel, 1985; Happel, 1980;Huang, 1990; Lee, 1987) in the 1970s is shown in FIG. 1. It showssignificant improvements over the conventional methanation and combinedshift/methanation processes. The sulfur-tolerant methanation processdeveloped by the Gas Research Institute (GRI) uses a molybdenum based(MoS₂) sulfur-tolerant catalyst (Happel, 1982). According to the GRIstudy, the process shows potential savings in steam usage, reducedrecycle rate and a smaller acid gas removal unit. This process waspiloted extensively by GRI from 1978 to 1985 in an adiabatic reactorsystem. The reactor was made from 1″ Schedule 80 pipe loaded 4″ deepwith ⅛″ cylindrical catalyst pellets. Their GRI-C-525 catalyst ran for10,000 hours and the GRI-C-600 catalyst ran for 2,300 hours.

When a nickel (Ni) catalyst is used for methanation and Ni catalyst issusceptible to coke formation, sulfur poisoning, and sintering. To solvethese problems, a couple of unit operations such as water-gas-shift(WGS) and acid gas removal (AGR) are required before the methanationreactor. Synthesis gas from the gasifier goes through a sulfur-tolerantWGS reactor to adjust the H₂/CO ratio to 3:1 (Reaction 1). Then, CO2 andsulfur compounds are removed from the hydrogen rich synthesis gas in anAGR unit before supplying it to a methanation reactor (Reaction 2).

Water-Gas-Shift Reaction:

CO+H₂O

CO₂+H₂(Exothermic)   (1)

Methanation Reaction:

CO+3H₂→CH₄+H₂O (Exothermic)   (2)

Catalytic steam gasification is another methanation technology, whichwas first piloted by Exxon (Nahas, 1983; Nahas, 2003; Nahas, 2004;Nahas, 1978; Nahas, 1978) in the 1970s (Anand, 2008). Recently,GreatPoint Energy (GPE), a new technology company, has done pilot plantcampaigns in a 1.5 ft reactor on catalytic steam gasification and hasplans to build a pilot facility in Somerset, Mass. Although iteliminates the need for an air separation plant, reduces the size of theacid gas removal unit and also combines gasification, shift andmethanation reactions into a single reactor, many operational issuesmust be proven in the Somerset pilot runs to determine if this processis economical.

One possible alternative source of hydrocarbons for producing fuels andchemicals is the natural carbon found in plants and animals, such as forexample, in the form of carbohydrates. These so-called “natural” carbonresources (or renewable hydrocarbons) are widely available, and remain atarget alternative source for the production of hydrocarbons. Forexample, it is known that carbohydrates and other sugar-based feedstockscan be used to produce ethanol, which has been used in gasohol and otherenergy applications. However, the use of ethanol in transportation fuelshas not proven to be cost effective and may not be achievable on a scalesignificant to current fuel requirements.

Carbohydrates, however, can also be used to produce fuel rangehydrocarbons. Although some upgrading technology has been developed toturn biologically derived materials into useful fuel and chemicalfeedstocks. Unfortunately, many carbohydrates (e.g., starches) areundesirable as feedstocks due to the costs associated with convertingthem to a usable form. In addition, many carbohydrates are known to be“difficult” to convert due to their chemical structure, the hydrocarbonproduct produced is undesirable, or the conversion process results inrelatively low yields of desirable products. Among the compounds thatare difficult to convert include compounds with low effective hydrogento carbon ratios, including carbohydrates such as starches, sugars,carboxylic acids and anhydrides, lower glycols, glycerin and otherpolyols and short chain aldehydes.

Cortright et al. US2008/0216391 teaches processes and reactor systemsprovided for the conversion of oxygenated hydrocarbons to hydrocarbons,ketones and alcohols useful as liquid fuels, such as gasoline, jet fuelor diesel fuel, and industrial chemicals. Abhari, US2009/0054701A1relates to a process for converting by products of the manufacture ofbiodiesel into industrially useful oxygenated products of greatercommercial value.

Molybdenum disulfide (MoS₂) based catalysts are important industrialcatalysts used in the removal of sulfur compounds from crude petroleumby hydrogenolysis. from Chemical Engineering nanostructured crystals ofMoS₂ (CE, April p. 15), a highly porous form of MoS₂ has also beenproduced by researchers at the University of Illinois atUrbana-Champaign (edlinks.che.com/4819-541). In U.S. Pat. No. 7,435,760,Hertling et al. propose using an alkali doped Cu catalyst, MoS₂catalyst, and Rh based catalysts to convert synthesis gas to higheralcohols. Different procedures and catalysts have been developed to getactive MoS₂ catalytic sites on the surface of the catalyst (Fujikawa,U.S. Pat. No. 7,361,624). Additionally, NiMoS and CoMoS catalysts arecommercially available from CENTINEL® and ASCENT® based on surfacedeposition of flat catalyst active sites, but this method of synthesisis limited to surfaces or platelets of catalyst activity which is notideal for all MoS₂ catalyzed reactions.

As such, development of an improved catalysts for convertingcarbohydrates, including “difficult” to convert starches as mentionedabove, to hydrocarbon, would be a significant contribution to the arts.In addition, development of a process for converting carbohydrates tohydrocarbons which yields significant quantities of desirablehydrocarbon products such as aromatics and olefins would be asignificant contribution to the art.

BRIEF DESCRIPTION OF THE DISCLOSURE

In one embodiment, a sulfur-tolerant methanation catalyst wassynthesized and a sulfur-tolerant methanation process was developed. Astable sulfur-tolerant methanation will increase catalyst longevity,increase product production and reduce plant operating costs. In oneembodiment, an MoS₂ catalyst with Zirconium promoter and elementalsulfur unexpectedly improved space velocity, temperature, and H₂Stolerance of the catalyst.

To convert biomass to liquid hydrocarbon fuels, a novel molybdenumdisulfide catalyst has been developed that improves the overall numberof active MoS₂ sites on the surface of the catalyst. Methods ofsynthesizing the active catalyst and use of the hydrogenation catalystare described. The ZrMoS hydrogenation catalyst converts C6 and C5sugars to sugar alcohols on a commercial scale. The ZrMoS catalystremains active in the presence of sugar alcohols and other products inaqueous biomass conversion process. Sugar alcohols, such as sorbitol,are hydrogenated to C6 hydrocarbons using a newly developed ZrMoScatalyst with a high density of surface active MoS₂ catalytic sites.

Molybdenum based sulfur-tolerant catalysts were studied to replacenickel catalysts for the conversion of synthesis gas to substitutenatural gas (SNG). Using a zirconium promoter with elemental sulfurduring co-precipitation synthesis increased the stability and activityof the catalyst, thus decreasing catalyst cost, increasing productivityand conversion rates. Other promoters and presulfiding conditions wereunable to provide these significant improvements. Process conditionssuch as space velocity, temperature, and H₂5 concentration increasedactivity of the MoS₂ catalyst.

A methanation catalyst comprising: molybdenum sulfate (MoS2), zirconia(Zr), and elemental sulfur (Sx), wherein the MoS2 methanation catalystis co-precipitated in the presence of Zr at a pH of greater than 3.0.

A process for hydrogenation of oxygenates comprising: contacting aZr/MoS2 methanation catalyst with R—C—O in the presence of hydrogen, andpurifying produced hydrocarbons, wherein the MoS2 methanation catalystis co-precipitated in the presence of Zr at a pH of greater than 3.0.

A substitute natural gas production system comprising: a gasifier unit,an oil quench unit, a water quench and ammonia recovery unit, amethanation unit, wherein the methanation unit comprises an MoS₂methanation catalyst, a CO₂ and water removal unit, and a sulfur removalunit, wherein said MoS2 methanation catalyst was co-precipitated in thepresence of Zr at a pH of greater than 3.0.

The methanation catalyst may be a co-precipitation of a MoS₂ catalystwith Zr at a pH of about 3.0, 3.5, 4.0, 4.5, 5.0, 6.0 or greater. Themethanation catalyst can use NiNO₃, MgNO₃, K₂CO₃, PdNO₃, Silica, orAlumina as promoters, as well as combinations of these promoters. Thecatalyst may be presulfided between ˜450° C. and ˜500° C., with N₂, H₂,DMDS, and combinations thereof; including but not limited to ˜500° C.with N₂, H₂ and DMDS; ˜500° C. with H₂ and DMDS; ˜500° C. with N₂, H₂and H₂S; or ˜500° C. with H₂ and H₂S. The catalyst may contain a ratioof Zr/Mo between about 0.6 and about 0.8. Additionally, the catalyst maybe synthesized with elemental sulfur.

BRIEF DESCRIPTION OF THE DRAWINGS

A more complete understanding of the present invention and benefitsthereof may be acquired by referring to the follow description taken inconjunction with the accompanying drawings in which:

FIG. 1: Unit operations required for the production of substitutenatural gas from synthesis gas using a nickel catalyst.

FIG. 2: Unit operations required for the production of substitutenatural gas from synthesis gas using a MoS₂ catalyst.

FIG. 3: CO conversion for methanation using MoS₂ catalysts prepared atdifferent pH levels during co-precipitation. Constant pH levels wereobtained by controlling the flow rates of the two peristaltic pumps forthe mixed salt solution and the 0.1M nitric acid solution.

FIG. 4: CO conversion and Zr/Mo ratio precipitated as a function of pHlevels during co-precipitation.

FIG. 5: CO conversion for methanation as a function of ratio of moles ofzirconium and molybdenum in order to obtain the optimized ratio forenhanced MoS2 catalyst activity.

FIG. 6: CO conversion and methane yield for sulfur-tolerant MoS2catalyst with or without zirconia as a promoter for the production ofsubstitute natural gas from synthesis gas.

FIG. 7: CO conversion for sulfur-tolerant MoS2 catalyst with or withoutaddition of elemental sulfur during the catalyst synthesis. Variousmultiples of the optimized amount of elemental sulfur (38 wt % of thecatalyst) were evaluated.

FIG. 8: (A) CO conversion and methane yield for sulfur-tolerant MoS2catalyst in the presence of 2200 ppm H2S obtained from the thermaldecomposition of dimethyl disulfide (DMDS). (B) CO conversion andmethane selectivity for sulfur-tolerant MoS2 catalyst in the presence of1% H2S obtained directly from a gas cylinder of 5% H2S in hydrogen.

FIG. 9: CO conversion and methane selectivity for sulfur-tolerant MoS2catalyst as a function of space velocity during the production ofsubstitute natural gas from synthesis gas.

FIG. 10: CO conversion for sulfur-tolerant MoS2 catalyst as a functionof H2S concentration in the feed gas for the production of substitutenatural gas from synthesis gas.

FIG. 11: CO conversion for sulfur-tolerant MoS2 catalyst as a functionof H2S concentration in the feed gas.

FIG. 12: CO conversion for sulfur-tolerant MoS2 catalyst as a functionof temperature during the production of substitute natural gas fromsynthesis gas.

DETAILED DESCRIPTION OF EMBODIMENTS OF THE INVENTION

Turning now to the detailed description of the preferred arrangement orarrangements of the present invention, it should be understood that theinventive features and concepts may be manifested in other arrangementsand that the scope of the invention is not limited to the embodimentsdescribed or illustrated. The scope of the invention is intended only tobe limited by the scope of the claims that follow.

Cellulose and hemicellulose are two major constituents in the biomassand can be broken down to C6 and C5 sugars using an acid or enzymehydrolysis process. C6 and C5 sugars can be further converted to sugaralcohols or other derivatives. The sugars and their derivatives can beupgraded to gasoline range hydrocarbons, mainly aromatics, using a ZSM-5catalyst, hydrotreating or combinations of ZSM-5 and hydrotreating.However, the sugars and sugar derivatives with less effective hydrogento carbon ratio are easily converted to coke and frequently lower liquidyield, foul expensive refining catalysts and other equipment. Additionof hydrogen donors with high effective hydrogen to carbon ratio such asmethanol (US4503278) and i-pentane (U.S. Pat. No. 7,678,950) have beenused to decrease coking, incorporated by reference. U.S. Pat. No.6,090,990 describes an improved catalyst containing a mixture of zeoliteand a binder treated with boron trichloride which is then used in theconversion of hydrocarbons to ethylene, propylene and BTX (benzene,toluene, xylene and ethylbenzene) aromatics. U.S. Pat. No. 7,550,634describes hydrotreating triglycerides to fuel range hydrocarbons. InU.S. Ser. No. 61/236,347, by Sughrue, et al., describes hydrotreating amixture of sorbitol and diesel over a commercial hydrotreating catalystto produce lighter alkanes and hexanes desirable for gasoline fuels.Additionally, in U.S. Ser. No. 61/248099, Yao, et al., describe theprocess of converting carbohydrates to gasoline boiling rangehydrocarbons by converting a carbohydrate-containing material to ahydrogenated carbohydrate material over a bi-functional catalyst andthen converting the hydrogenated carbohydrate material to gasolineboiling range hydrocarbons over a zeolite catalyst. In U.S. Ser. No.61/288,912, Yao, et al., use a zinc-platinum or cobalt-molybdenumimpregnated zeolite catalyst (ZnPt-zeolite or CoMo-zeolite) with acarbohydrate or polyol to produce polyols and hydrocarbons. In U.S. Ser.No. 61/424,896, Bares, et al., use a single-step hydrotreating processto convert oxygen-containing hydrocarbons (preferably, biomass-derivedhydrocarbons) that allows a lower conversion temperature to be utilizedrelative to conventional hydrotreating over a CoMo catalyst. Thesepatents and applications referenced above are specifically incorporatedby reference in their entirety.

Carbohydrates, such as starches and sugars may be converted inaccordance with the present invention to form a hydrocarbon mixtureuseful for liquid fuels and chemicals. The term, “carbohydrate” is usedgenerally to refer to a compound of carbon, hydrogen and oxygen havingthe general formula C_(x)(H₂O)_(y), in which the ratio of hydrogen tooxygen is the same as in water. Carbohydrates include monosaccharides,polysaccharides, and mixtures of monosaccharides and/or polysaccharides.The term “monosaccharide” or “monosaccharides” are generally hydroxyaldehydes or hydroxy ketones which cannot be hydrolyzed into any simplercarbohydrate. Monosaccharides can be a triose with 3 carbon atoms,tetrose with 4 carbon atoms, pentose with 5 carbon atoms, hexose with 6carbon atoms, or larger monosaccharides like Sedoheptulose with 7 carbonatoms or Neuraminic acid with 9 carbon atoms. Examples ofmonosaccharides include glyceraldehyde, erythrose, xylose, dextrose,glucose, fructose and galactose. The term “polysaccharide” or“polysaccharides” include those saccharides containing more than onemonosaccharide unit. This term also includes disaccharides (such assucrose, maltose, cellobiose, and lactose) and oligosaccharides.

Carbohydrate feedstock comprises a mixture of one or more carbohydratederivatives including polysaccharides, monosaccharides, polyols, sugarsand sugar alcohols from a variety of sources, as well as otherbyproducts of biological degradation that aren't removed as solids orare not completely removed by other processes. In some examples a singlepolyol, such as sorbitol or xylitol in aqueous solution is used as acarbon feedstock. Sugar feedstocks consist of one or more polyols in anaqueous solution. Polyols include glycerol, sorbitol, xylitol, and thelike. Liquefaction of biomass typically produces monoglyceridefeedstocks containing sorbitol and xylitol. Feedstocks may contain fromabout 50 to about 98% v/v polyol. In one embodiment a polyol feedstockcontains approximately 30%, 35%, 40%, 45%, 50%, 55%, 60%, 65%, 70%, 75%,80%, 85%, 90%, 95%, up to 98% sorbitol, xylitol and mixtures of sorbitoland xylitol. Although sorbitol feedstock comprises sorbitol and aqueoussolution, additional polyols, oils, and sugars are present afterliquefaction. Many isomers, polymers, and soluble sugars are present inthe aqueous liquefaction fraction. Examples of carbohydrates useful asstarting materials in accordance with the present invention include, butare not limited to, polysaccharides such as sucrose, maltose, lactose,cellobiose, melibiose, raffinose, starch (derived from a variety ofcereal grains such as wheat and rice, tubers such as potato, tapioca,and arrowroot, or waxy starches such as waxy maize) and starchdecomposition products such as dextrin and corn syrup (also known asglucose syrup).

Sulfur-Tolerant Methanation: Molybdenum catalyst has both WGS andmethanation activity. The reaction occurring in the methanation reactoris likely to be a combination of reactions 1 and 2 (above) as shown inreaction 3. This eliminates the need for a WGS reactor.

Methanation Reaction:

2CO+2H₂→CO₂+CH₄ (Exothermic)   (3)

Apart from being sulfur-tolerant, MoS₂ catalyst also methanates the rawsyngas directly by using equimolar amounts of carbon monoxide andhydrogen. There is no requirement for a water-gas-shift reactor beforethe methanation reactor with a MoS₂ catalyst, resulting in steam usagesavings. Also because the MoS₂ catalysts are sulfur-tolerant, sulfurimpurities have to be removed to only 4 ppm levels in the product gas(pipeline specification) instead of 20 ppb in the synthesis gas (nickelcatalyst specification). In addition, the position of the acid gasremoval unit can be changed from upstream of the methanation reactor todownstream of the methanation reactor resulting in a decrease in gaseousmoles processed by the acid gas removal unit. This results in potentialsavings due to the smaller size of acid gas removal unit. Molybdenumcatalysts also operate at higher temperature than Ni catalysts resultingin lower catalyst sintering during methanation.

EXAMPLE 1 Promoter Effects

MoS₂ catalysts with various promoters were prepared in this study.Catalysts were prepared by co-precipitation using ammoniumtetrathiomolybdate as the molybdenum precursor. Catalyst preparationbegan with the precipitation of amorphous MoS₃ along with ZrO₂. Variousother promoters were also tested along with zirconium. A salt solutionwas prepared by mixing two aqueous solutions—4 grams of ammoniumtetrathiomolybdate in 61.52 mL of water, and 1.776 grams of zirconylhydrate nitrate in 6 mL of water. Two peristaltic pumps were used toslowly add dilute nitric acid (0.1 M) and the mixed salt solution to a500 mL beaker filled with 100 mL water. The beaker was vigorouslystirred and pH was maintained constant throughout the precipitationprocess by adjusting the flow rates of the solutions. Precipitation wasdone at various pH levels to study their impact on the activity of thecatalyst. The resulting slurry was filtered and washed with distilledwater and acetone. The filtered cake was added to 1.538 grams of sulfurstirred in acetone. The acetone was allowed to evaporate and then theresultant cake was further dried in an oven under N₂ atmosphere for 4hrs at 80° C. (˜175° F.). The dried mixture was collected, weighed,pelletized and prepared for presulfiding and reduction for furtherevaluation.

TABLE 1 CO conversion and methane selectivity for various promoters COCH₄ Conversion Selectivity Deactivation Promoters (%) (%) Rate (%/hr)MoS₂ Catalyst with Zirconia 85 52 0.005 Promoters 1% NiNO₃ 76 51 0.25510% NiNO₃ 36 46 0.256 10% MgNO₃ 85 52 0.083 3.5% K₂CO₃ 83 52 0.008 1%PdNO3 67 51 0.471 1% Silica 80 52 0.026 1% Alumina 80 52 0.013

Other promoters such as Ni, Pd, Mg, K, Al, Si, and Ti were added at 1%by catalyst weight with or without zirconium during co-precipitation.Catalytic materials were characterized by Analytical Services usingx-ray fluorescence (XRF), nitrogen physisorption (BET) and x-raydiffraction (XRD) to determine elemental content, BET surface area andcrystal structure.

To evaluate the catalyst activity, 3.8 mL of MoS₂ catalyst mixed with6.2 mL of alundum was loaded into a ½ inch stainless steel reactor toproduce a 10 mL catalyst bed. The catalyst was pre-sulfided by rapidlyheating the reactor at 460 psig with N₂ flowing at a rate of 45 sccmalong with 3% H₂S. H₂S can be provided either directly as H₂S gas or bythe thermal decomposition of dimethyl disulfide liquid (DMDS). Thereactor was heated to about 500° C. (˜930° F.) in less than 15 minutesto have fast reduction and to obtain a high surface area MoS₂ catalyst.After achieving 450° C. (˜840° F.), the nitrogen gas flow was switchedto a stream of hydrogen (45.5 sccm) along with DMDS flow at 0.15 mL/hr.The catalyst was held under H₂ and DMDS flow at 500° C. for 5 hrs tofully reduce the MoS₃ phase to the MoS₂ phase. After 5 hours, thecatalyst was evaluated using reaction conditions.

DMDS feed as a H₂S source was pumped to the system using an ISCO Model500 D syringe pump while gases were supplied by Brooks 5850E mass flowcontrollers. The temperature at the center of the catalyst bed wasmeasured using a type K thermocouple inside a thermowell in the reactor.Pressure was maintained using a Tescom back pressure regulator. Analysisof the reactor effluent was completed using an online Agilent 6890 gaschromatograph outfitted with a 15′×⅛″ stainless steel 60/80 mesh sizecarboxen-1000 column (0.5 g/ft packing density) plumbed to a thermalconductivity detector. Periodically a gas bomb was used to collect thereactor outlet, which was analyzed either by detailed hydrocarbonanalysis or by mass spectroscopy.

A series of MoS₂ catalysts were prepared at various pH levels byco-precipitation to study the effect of elemental sulfur, types andcomposition of various promoters, presulfiding chemical and presulfidingtemperature. These catalysts were evaluated at a feed composition (mol)of 34% CO, 37% H₂ and 28% N₂. A H₂/CO mole ratio of 1.08 was picked formolybdenum catalysts because the E-Gas gasifier can provide this ratiowith some types of coals by putting some extra steam in the second stageof gasifier. A space velocity of 2400 hr⁻¹ was used because it is closeto the fresh feed space velocity for conventional methanation (Nicatalyst) in industrial plants, which is around 2000 hr⁻¹. Catalystswere evaluated at various percentages of H₂S, space velocities, andreaction temperatures.

For the activity test, unless otherwise stated, reactor pressure was 460psi with a gas hourly space velocity (GHSV) of 2400 hr⁻¹ and hydrogen toCO ratio (H₂/CO) of 1.08. Furnace temperature was set at 455° C. (−850°F.). The syngas feed stream was 37% H₂, 34% CO, 1% H₂S and the remaininginerts in the syngas were substituted with N₂ flow at 28%. Thesereaction conditions were chosen to compare the catalyst activity withthe GRI catalyst.

EXAMPLE 2 Catalyst Synthesis Parameters:

Effect of pH during Precipitation: In this work, a constant pH of 3, 4,5, or 6 was maintained during catalyst synthesis. The constant pH wasobtained by using two peristaltic pumps set at appropriate flow rates tomix the salt solution with the nitric acid solution (0.1M). Catalystsobtained were evaluated and compared for CO conversion and methaneselectivity. Selectivity of methane was almost the same for all thecatalysts prepared. The effect of pH on CO conversion is shown in FIG.3. It is clear that the catalyst precipitated at pH 5 gives the maximumconversion. After analyzing the catalyst samples with PVSA (pore volumesurface area), Karnak and XRD studies, it was observed that the effectof pH on CO conversion is not due to the presence of different catalystphases or surface areas but due to the ratio of Zr to Mo precipitated inthe catalyst synthesis. As shown in FIG. 4, the Zr/Mo ratio (mol)precipitated during catalyst preparation follows a trend with the pH ofthe solution. Near pH 5 an optimized ratio of Zr to Mo was precipitatedcompared to ratios precipitated at other pH levels. The optimum ratio ofZr/Mo lies between 0.6 and 0.8. Because it is hard to make very smallchanges in the pH in order to find exact optimum Zr/Mo ratio,experiments were run to find out the optimum ratio of zirconium tomolybdenum in the bulk salt solution.

EXAMPLE 3 Zirconium Promoter

Experiments were run to find out the optimum amount of zirconium in thebulk salt solution for the given amount of molybdenum for the improvedactivity. It can be noticed from the plot (FIG. 5) between CO conversionand the ratio of moles of Zr to Mo, that there is a maximum in the COconversion at a ratio of 0.75 moles of zirconium to molybdenum. Up to aratio of 0.75, there is an increase in the CO conversion and itdecreases with the further increase in zirconium amount. FIG. 6 showsthe CO conversion and methane yield data for MoS₂ catalysts with andwithout zirconia as a promoter. It is clear from these data thatzirconia enhanced the catalyst activity without changing the stabilityof the catalyst under the given conditions. Zirconia improved the COconversion from 73% to 85% and methane yield on the basis of weight from21.3% to 23.7%.

The basic structure of the catalyst can be varied by the incorporationof additional elements such as nickel, magnesium, potassium, palladium,aluminum, silicon, or titanium as promoters (Cover, 1989). In thisstudy, catalysts were synthesized with these elements at 1-10 wt % ofthe catalyst. As shown in the Table 1, CO conversion and methaneselectivity did not show any improvement compared to MoS₂ catalyst with0.75 moles of zirconium to molybdenum. There was also not much influenceon the stability of the catalyst as it is shown from the deactivationrate. Deactivation rate is given in terms of loss in percentage of COconversion per day.

EXAMPLE 4 Elemental Sulfur

It was discovered that addition of elemental sulfur during theco-precipitation of the MoS₂ catalyst, followed by its removal as H₂S byreduction, improves catalytic stability and activity. MoS₂ catalystswere synthesized with and without elemental sulfur and analyzed formethanation activity. It is clear from FIG. 7 that sulfur is importantfor stability and its quantity has a small effect on the catalystactivity. The black curve is shown for CO conversion for a catalystwithout sulfur. There is decline over time in the CO conversion from 84%to 79% as seen from 45 hours of data for the MoS₂ catalyst synthesizedwithout sulfur. On the other hand, the dark curve shown for theoptimized amount of sulfur (38 wt % of the catalyst) is stable and COconversion is constant at 84% for 45 hours. CO conversion for the redcurve, shown for half the optimized amount (20 wt % of the catalyst),also declines from 81% to 79%. Alternatively, the green curve, shown fordouble the amount of optimized sulfur (77 wt % of the catalyst), isfairly stable but CO conversion is around 76%.

EXAMPLE 5 Presulfiding Conditions

Effect of Presulfiding Temperature: The surface area of molybdenumdisulfide depends strongly on the rate of conversion of molybdenumtrisulfide to disulfide. Rapid conversion of trisulfide to disulfide at450° C. either by reduction with hydrogen or thermal decompositionresults in molybdenum disulfide of unusually high surface area: 135-155m2/gm. On the other hand, slow conversion gives surface areas as low as2 m2/gm. There is an optimum temperature around 450° C. for carrying outreduction. This optimum temperature is due to competition between therate of nucleation of MoS₂ and rate of sintering of MoS₂ crystals afterformation. A number of experiments were conducted to evaluate differentreduction temperatures around 450° C. and results are given in Table 2.Since the CO conversion was decreasing with temperature, the MoS₂catalyst was not tested below 450° C.

TABLE 2 Comparison of CO conversion and methane Selectivity for variouspresulfiding temperatures and presulfiding media evaluated during MoS2catalyst reduction and presulfiding CO Conversion CH₄ Selectivity (%)(%) Presulfiding Conditions Presulfided at 500° C. with N₂, 85 52 H₂ andDMDS (Base Case) Presulfiding Temperature 500° C. 85 52 480° C. 83 52450° C. 82 52 Presulfiding Medium N₂, H₂ with DMDS 85 52 H₂ with DMDS 7952 N₂, H₂ with H₂S 81 52 H₂ with H₂S 85 52

Effect of Presulfiding Medium: The MoS₂ catalyst was evaluated fordifferent presulfiding media: hydrogen sulfide, a gaseous medium, anddimethyl disulfide, a liquid sulfiding medium in the presence of gasessuch as N₂ and H₂. Molybdenum trisulfide can be converted to disulfideeither by reduction with hydrogen or thermal decomposition in N₂. Anumber of experiments were done to reduce and presulfide the MoS₂catalyst with the two different sulfiding agents, as shown in Table 2.No major differences in MoS₂ activity for methanation were observed fordifferent presulfiding media.

Results:

A 2.5 day run was conducted on the MoS₂ catalyst with 2200 ppm H₂Sobtained from thermal decomposition of dimethyl disulfide (DMDS). A plotof CO conversion is shown in FIG. 8( a). The catalyst resultant in 95%conversion, but due to the low flow of the DMDS syringe pump, there weresome scattering in the data. In order to obtain stable data without thedisturbances, a gas cylinder of 5% H₂S in H₂ was bought and an 11-dayexperiment was conducted with feed gases containing 1% H₂S. The resultwas stable after some initial deactivation as shown in FIG. 8( b). TheCO conversion obtained was around 80%. It was higher than previous COconversion of 76% (Meyer, 1982) for 1% H₂S. The catalyst was evaluatedwith feed gases simulating raw gasifier effluents for extended periodsto measure the effects of CO₂, H₂O, benzene, phenol, ammonia, and higherhydrocarbons, etc.

Effect of Process Conditions: The effects of various parameters such asH₂S in the synthesis gas, space velocity, temperature and pressure onthe activity and stability of the catalyst influence catalyst selectionand design of the process.

Effect of Space Velocity: The effect of GHSV on CO conversion andmethane selectivity for the MoS₂ catalyst is presented in FIG. 9. Asexpected, increases in GHSV decrease CO conversion and this decrease wasnot linear. The CH4 selectivity was fairly constant over the entirerange of space velocity examined (2000 hr⁻¹ to 20,000 hr⁻¹). Since themethanation reaction is so exothermic, the space velocity should bechosen so that there is just enough conversion to keep the temperaturewithin the acceptable limits.

Effect of H₂S in Synthesis Gas: The effect of H₂S in the feed gas on COconversion was measured and is given in FIG. 10. As shown in FIG. 10,there is a sharp decrease in CO conversion with increasing H₂Sconcentration. However MoS₂ is a sulfur-tolerant catalyst and so theeffect of H₂S on this catalyst is reversible. FIG. 11 shows that whenH₂S concentration was increased from 1% to 2% of the feed gas, the COconversion reduced from 80% to 71%. But when the H₂S level was reducedback to 1%, the CO conversion returned to 80%. These data show that forMoS₂ catalysts, H₂S is more like an inhibitor than a poison. On theother hand, H₂S acts as a poison for the nickel catalyst because thereduction in the activity is not reversible.

Effect of Temperature: In order to check the upper temperature limit ofthe MoS₂ catalyst, catalysts were evaluated at temperatures higher than455° C. FIG. 12 shows the CO conversion for two furnace set pointtemperatures of 455° C. and 600° C. over a period of 11 days. Thesefurnace set point temperatures corresponded to catalyst bed temperaturesof 484 and 640° C. (˜1180° F.), respectively. It is clear from the plotthat initial conversion is 90% at a catalyst temperature of 640° C.compared to 89% at 484° C. But the decline in activity is much faster atthe higher temperature. The deactivation rate from the last two days ofdata from 11-day experiment is 0.009% loss in CO conversion per hr for484° C. (˜900° F.) and 0.045% loss in CO conversion per hr for 640° C.The higher deactivation rate at the higher temperature is due to moresintering. A deactivation study longer than 11 days would be required tosee if the catalyst finally stabilizes at higher temperatures.

EXAMPLE 6 Co-Precipitation in Acid

The MoS₂, catalyst along with zirconium as promoter was prepared byco-precipitation using nitric acid. 32 grams of ammoniumtetrathiomolybdate was dissolved in 492 mL of distilled water. 14.2grams of zirconyl hydrate nitrate dissolved in 30.2 grams of DI waterwas added to this solution. The solution was acidified with 0.1 M dilutenitric acid at pH 5 to precipitate the MoS₂ catalyst. The precipitatewas then filtered, washed first with distilled water and then acetone.Then 40 grams of elemental sulfur stirred in acetone was added to thefiltered cake. The resultant cake was dried in an oven under N₂atmosphere at 80° C. for 4 hrs. The dried mixture is collected, weighed,palletized and prepared for presulfiding and reduction for furtherevaluation.

To evaluate the catalyst activity, 15 g of MoS₂ catalyst were loadedinto a ½ inch stainless steel reactor and pre-sulfided by rapidlyheating the reactor to 450° C. at 460 psig with N₂ flowing at a rate of187.5 cc/min along with DMDS flow at 0.4 ml/hr. When the temperaturereached 450° C., the nitrogen gas flow was switched to a stream ofhydrogen along with DMDS flow at 0.59 mL/hr. The reactor was kept at500° C. under H₂ and DMDS for 5 hours. After 5 hours, the reactor wascooled by passing a stream of N₂ through the reactor and used forfurther evaluation. For the activity test, sorbitol was selected as C6sugar alcohol. Diesel was co-fed to the reactor as diluent.

The detailed reaction conditions and catalyst performance are listed inTable 3, below. It is clearly shown that the catalyst is active forsorbitol hydrogenation to hydrocarbons. The sorbitol conversion wasabout 81% at 250° C. (˜480° F.). With the increase in reactiontemperature to 315° C. (˜600° F.), the sorbitol conversion was increasedto ˜95%. At 280° C. (˜535° F.) the sorbitol conversion was approximately91%. C6 hydrocarbons, such as hexanes, are the main products producesfrom sorbitol conversion.

TABLE 3 Sorbitol Conversion to C6 hydrocarbons at varying temperaturesRun# 250° C. 280° C. 315° C. Temp, C. 250 280 315 Pressure, Psig 12001200 1200 70% sorbitol rate, ml/min 6 6 6 Diesel rate, ml/min 12 12 12H₂ rate, ml/min 200 200 200 Sorbitol conversion, % 81.0 91.2 95.0 C6hydrocarbons in collected liquid product, 10.6 19.0 19.7 wt %

In conclusion, CO conversion was increased from 75% to 85% (455° C., 460psig, GHSV=2400 hr⁻¹, H₂S=1% (mol), H₂/CO=1.08) by using a zirconiumpromoter composition during co-precipitation of a high surface area MoS₂catalyst at the appropriate pH. Elemental sulfur powder is important forthe stability of the MoS₂ catalyst. Addition of promoters such asnickel, magnesium, potassium, palladium, alumina, silica and titania didnot result in any improvements in the methanation activity. High surfacearea MoS₂ catalyst is resultant from fast reduction of molybdenumtrisulfide in hydrogen into the active molybdenum disulfide form. Thechoice of presulfiding agent, dimethyl disulfide or hydrogen sulfide,did not result in differences in methanation activity. Presence of highlevels of hydrogen sulfide reduces the methanation activity of the MoS₂catalyst but activity returns when the hydrogen sulfide level drops.Molybdenum catalyst was tested at 640° C. and it retained activityalthough deactivation rate was 0.045% compared to 0.009% loss in COconversion per hour at 484° C. Despite substantial improvement,molybdenum catalysts are less active than nickel catalysts formethanation. However, very high activity is not required because of highheat generation in methanation which places practical limits on per-passconversion. Thus the novel sulfur tolerant MoS2 catalyst provides acatalyst that is active at higher temperatures allowing uniqueMethanation process design and implementation due to unique catalyticproperties.

REFERENCES

All of the references cited herein are expressly incorporated byreference. The discussion of any reference is not an admission that itis prior art to the present invention, especially any reference that mayhave a publication data after the priority date of this application.Incorporated references are listed again here for convenience:

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1. A substitute natural gas production system comprising: a gasifierunit, an oil quench unit, a water quench and ammonia recovery unit, amethanation unit that comprises a methanation catalyst, a CO₂ and waterremoval unit, and a sulfur removal unit, wherein said methanationcatalyst comprises molybdenum disulfide (MoS₂), zirconium, (Zr) andelemental sulfur, and wherein the methanation catalyst isco-precipitated in the presence of Zr at a pH of greater than 3.0. 2.The system of claim 1, wherein said methanation catalyst isco-precipitated in the presence of Zr at a pH of about 3.5, 4.0, 4.5,5.0, 6.0 or greater.
 3. The system of claim 1, wherein said methanationcatalyst comprises a promoter that is selected from the group consistingof NiNO₃, MgNO₃, K₂CO₃, PdNO₃, silica, alumina, and combinationsthereof.
 4. The system of claim 1, wherein said methanation catalyst ispre-sulfided at between about 450° C. and about 500° C. in the presenceof one or more members of the group consisting of N₂, H₂, DMDS and H₂S.5. The system of claim 1, wherein said methanation catalyst contains aratio of Zr/Mo between about 0.6 and about 0.8.